Mixed-phase hydrofining of hydrocarbon oils



Sept 13, 1960 A. E. KELLEY Erm.

MIXED-PHASE HYDROFINING oFL HYDRocARBoN oILs Filed Aug. '5, 1 95'? nited States Patent O MIXED-PHASE HYDROFINING OF HYDRO- CARBON OILS Arnold E. Kelley, Fullerton, and Roland F. Deering, 'Whittien Calif., assignors to Union Oil Company of California, Los Angeles, Calif., a corporation of California Filed Aug. 5, 1957, Ser. No. 676,058

12 Claims. (Cl. 20S-210) This invention relates to the catalytic hydroning of mineral oils wherein the conditions of treatment are such that a portion of the feed normally remains in the liquid phase, and another portion remains in the vapor phase during processing. More specically, the invention is concerned with means for increasing the degree of hydroiining of the vapor-phase fraction relative to that of the heavy fraction, but without necessarily increasing the severity of the processing with respect to the liquid-phase fraction.

The gist of the invention consists in first subjecting the total feed to mixed-phase hydroning in a first catalytic hydroiining zone until the liquid phase is sufciently refined, then eifecting a separation of the gas phase from the liquid phase, said separation being performed at essentially the pressure prevailing in the first reactor, and transferring the separated gas phase to a second catalytic hydroning Zone for further treatment to remove more completely the remaining sulfur compounds and nitrogen compounds. In a preferred aspect of the invention, fresh makeup hydrogen for the system is lirst employed as a stripping gas for the liquid-phase product from the first hydroiining zone, and is then mingled with the vapor-phase products from the rst hydroiining zone, and the combined gas phase is then treated in the second hydrotining zone. Recycle hydrogen from the second hydroning zone is then recycled to the first hydroiining zone. In still another modification, the total effluent from the iirst hydroning zone is partially cooled prior to separating the gas phase from the liquid phase. In this manner part of the gaseous products are condensed, and control may be maintained over the boiling range of the vapor-phase product which is to be further treated in the second hydrofining zone.

One of the principal objects of the invention is to provide simple and economical methods for the controlled catalytic hydroiining of wide-boiling-range feedstock. A more specific object is to provide convenient and inexpensive methods for subjecting a selected light fraction of a given feed to a more severe hydroning treatment than the remaining heavy fraction. Another object is to provide optimum catalysts and conditions for liquid-phase and for vapor-phase hydroiining. Another object is to provide for the most eicient possible utilization of catalyst where it is desired to hydrone `rigorously the light ends of a wide-boiling-range feedstock, and to hydroiine lightly the heavy ends thereof. Another object is to provide novel combinations of apparatus for effecting the herein-described hydrofining treatments. Other objects and advantages will be apparent from the more detailed description which follows:

'I'he term hydroiining as used herein means the selective hydrocracking of hydrocarbon feedstocks contaminated with various organic impurities such as sulfur compounds, nitrogen compounds, and oxygen compounds, with resultant net chemical consumption of hydrogen. The catalysts used, and the reaction conditions are chosen ICC so as to effect a selective hydrogenation and decomposition of the sulfur, nitrogen, and oxygen compounds, without causing any appreciable hydrocracking of the hydrocarbon components. Such hydroining processes have become widely used for rening selected feedstocks, e.g. gasolines, heavy gas oils, light gas oils, kerosene, solvent naphthas, and the like. In these known processes, the feed is admixed with e.g. 30C-5000 s.c.f. of hydrogen per barrel of feed, preheated to a temperature of about 650-875 F., and then passed through a bed of the desired catalyst. Pressures of about 300 to 5000 p.s.i.g. are normally employed, along with feed rates amounting to about 0.5 to 15 volumes of feed per volume of catalyst per hour. A

Under the foregoing processing conditions, the light feeds such as gasoline are normally present wholly, or almost wholly, in the vapor phase. Heavier materials such as gas oils are normally present both in the liquid and the vapor phase. The present invention is concerned specically with the problems which arise when such heavy, and/or wide-boiling-range feeds, are employed that a considerable proportion of the total feed, e.g. 20% to would normally be present in the vapor phase, the remainder going through as liquid.

Where a liquid phase is present, it Will form a liquid Iilm covering the entire active catalyst surface. Consequently, the feed which remains in the gas phase can react only by diffusion through the liquid lm on the catalyst. It is obvious that under such conditions, the liquid-phase portion of the feed will tend to react more rapidly than the vapor phase, at least under conditions Where the volume of liquid present, and its state of turbulence, are not such as to too greatly impair mass-transfer rates between liquid and catalyst. In any event, the extent of hyroiining of the liquid phase which occurs in any given pass through the reactor is seldom the same as that of the accompanying vapor-phase hydrocarbons.

However, the process of this invention is useful not only in those cases where the conversion rates of gas and liquid phases are different, but in all cases where the objective is to effect a more complete hydrofining of the light fraction than can be obtained without over-treatment of the heavy fraction in normal mixed-phase hydrolining. This condition may occur for example where it is desired to process a 300-650 F. boiling-range fraction which would include both heavy naphtha and light gas oil, the light gas oil being subsequently used for diesel fuel or cracking stock, and the heavy naphtha being used as reforming stock. For purposes of reforming over sulfursensitive catalysts such as platinum, it is necessary to reduce the sulfur content to below about 0.05% and the nitrogen content to below about 0.0001%. On the other hand, cracking charge stocks or diesel fuels can tolerate substantially larger amounts of sulfur and nitrogen compounds. Another instance of utility involves the treatment of a wide-range gas oil, wherein the light gas oil is to be used as jet fuel, and the heavy gas oil as cracking charge stock.

In still other instances, product specifications for the light and heavy fractions may be substantially the same in terms of sulfur and nitrogen content. However, where only a small part of the charge is in the liquid phase, e.g. 5-20%, it may become suiiiciently hydroiined before product specifications are met on the vapor-phase portion, due to the differential conversion rates discussed above. It would be disadvantageous to increase the overall severity of treatment to obtain the desired vaporphase conversion, because this would involve overtreatment of the liquid-phase portion with resultant losses in product yield and increased rates of catalyst deactivation resulting from gum and coke deposits. Hence, the twostage hydroning of this invention may be useful even where a single ultimate product is desired.

It will hence be apparent that the process of` invention is useful wherever it is desired'to hydroiinea` charge stock in a single treating unit under conditions where part of the charge isnormally liquid and part is gaseous, and wherein it is desired that a light fraction of the charge be-subjected to additional hydroning in the absenceof the heavy portion.

Feedstocks which maybe treated herein include in general Vany mineral oil stock having an end-boiling-point in excess of about 450Y F. and an initial boiling point at least about 75 `F. lower than -the end-boiling-point. When using stocks of this type, Yit isalmost inevitable that the'heavy ends will constitute a relativelyV fixed and unchanging liquid phase during hydrofming, while the light endsy will undergo treatment predominantly in ther-'vapor phase. Specific examples-of such stocks include crude oils,rreduced crude Voils, deasphalted reduced crude oils, lightgas oils, yheavy gas oils, kerosene-gas oil fractions, heavy naphtha-gas oil fractions, fuel oil fractions, etc. These stocks may be derived from petroleum, shale, tar sands and similar natural deposits. Y

Operation of the process may be more readily understood with reference `to the accompanying drawings. Figure l is a schematic flow diagram illustrating one modiiication, and Figure 2 is a schematic flow diagram illustrating a slightly diiferent modication. Neither of these illustrations however is intended to be limiting in scope.

InFigure l, the principal piece of apparatus comprises a reactor-stripper 1, comprising an enclosed upper cylindrical reactor shell 2, and a communicating, subjacent cylindricalstripping column `3. A bed of suitable granular hydroning catalyst 4 is supported in reactor shell 2 by means of a perforated supporting plate 5. Immediately below plate 5 is provided a substantial open space 6 wherein phase separation of the initial hydrofining products takes place. Y

Stripping'column 3 is removable'attached'to reactor shell 2, and is preferably of smaller diameter. The interior of stripping column 3 preferably contains a series of perforated plates 8, or bubble cap trays of conventional design. Alternatively7 stripping column Smay be packed with suitable conventional packing such as Raschig rings, glass beads, porcelain chips, etc. The entire reactorstripper 1 may suitably be constructed of mild steel, stainlessisteel, .or any other pressure-retaining, corrosion-resist-ant material which will withstand temperatures up to about 1000 'F.

In operation, the initial feed is brought in through line 10,.pump 12, heat exchanger 13, nal preheater 14, and line into the top of reactor 2. Recycle hydrogen from line 11 is mingled with the preheated feed in line 15. The mixture of feed and hydrogen then fiows downwardly through catalyst bed 4 and is discharged into separation zone 6. 7T he liquid portion of product falls by gravity into stripping column 3, and the vapor phase portion is withdrawn from the reactor through line 17. A segmental deector plate welded to the inner surface of reactor shell 2 serves to deect liquid product away from vapor outlet line 17 and preventsentrainment of liquid in the withdrawn gases.

The liquid phase accumulating in stripping column 3 percolates downwardly countercurrently Ito a stream of hydrogenV admitted via line 22 near the bottom of stripping column 3. This stripping hydrogen preferably cornprlses all or a part of the makeup hydrogen required in the process, Le. the hydrogen required to replace that which is chemically consumed in the process. Itmay be introduced at a low temperature of e.g. 50i-300 F., and becomeheated by contact with thev oil, or it may beintroduced at a higher temperature. The fresh hydrogen is added via line 16 at a rate regulated by valve 18 and 4 pressure controller 19 so as to maintain `the Vdesired pressure in the system.

'The use of fresh hydrogen as a stripping agent is desirable in that it provides for a more efcient stripping of the light fractions from the liquid infthe column. If additional gas is required for effective stripping, a portion of the nal recycle gases may be A:admixed with the fresh hydrogen, the latter being admitted via line 24 to recycle line-40, and the whole being injectedinto column 3 by means of pump 25. Thehydrogen introduced via line 22 then accumulates in separator space i6 and is withdrawn via line 17 along with the vapor-phase products from reactor ,2. The liquid product in column 3 is withdrawn vialine 27, the rate of'withdrawal being controlled by valve 28 operated by liquid-level controller 29.

The total Vapor-phase product in line 17 is then transferred directly to vapor-phase reactor 30, which is conventional in design` and consists of a cylindrical iron or steel shell containing a bed of suitable hydroning catalyst supported on perforated plate 31. Vapors pass downwardly through reactor 30 in contact with the catalyst to carry the hydroining to the desired degree of completion. The efliuent from vapor-phase reactor 30 is withdrawn via line 32, precooled inexcnanger 33 against incoming fresh feed or other cool process stream, and'subjected to final cooling in condenser 35. The condenser product then collects in accumulator 36, which is maintained at substantially the same pressure as reactor 30.' Recycle gases are removed from separator 36 via line V38, and passed either in whole or in part via line 39 and preheater 37 to mingle with fresh feed in line 15. A portion of the recyclegases maybe passed via line 40 for use in l stripping column 3 as previously described. A small portion of recycle gas maybe bled off via line 41 to prevent the buildup of inerts in the recycle stream.

The liquid product in separtor 36 is then transferred rto low-pressure separator '42 via line 43, and liquid-levelcontrolled valve 44. In low-pressure separator 4Z, offgases containingrmainly methane, ethane, hydrogen suliide and hydrogen are withdrawn via line 45. 'Ihe liquid product is withdrawn via line 46 and transferred to final fractionating units not shown. Y

Referring now to Figure 2, the general processing scheme is similar to that in Figure l but is modified to permit a greater degree of control over the boiling range of the liquid product withdrawn after initial reaction, -as well as the boiling range of the light fraction which is subjected to further processing. `ln this modification the initial feed is brought in via line 50, mingled with recycle hydrogen from line 51, and transferred via heat `exchanger 52, final preheater 53, and line 54, into the top of mixedphase reactor y55. The constructionA of the processing equipment and the instrumentation in this modilication is similar to that in Figure l and hence will ,not be described in detail. The mixed feedl passes through catalyst bed 56 and is withdrawn via line 58.

The product in line 58 is then passed through a heat exchanger -59 to effect partial cooling, i.e. to the degree required to condense the zdesired portion `of the vaporphase product. Thepartiallycondensed mixture is then transferred via line 60'to separtor-strripper 61. The total products are dashed into open space 63 lto allow separationiof liquid'by `deceleration and gravity. Theivapor phase from separator 63 `isY then withdrawn via line 65, reheated in heater 66 to the desired reaction temperature, and then admitted -to vapor-phase reactor 67 'wherein final processing takes place as previously described for reactor 30.y The 4vapor-phase product from reactor 67 is thenwithdrawn yialine 68, precooled in exchanger V69, condensed in condenser` 7 0, and thenV passed into highpressure separator 71. Recycle gas is withdrawn from separator 71 Vand Vtransferred via lines 73 and 51 to feed line 50, and may in part be diverted through line 75 to be `used as stripping gas for the liquid product in stripping column 77.

Stripping column 77 is operated similarly to column 3 of Figure 1. Here yagain it is preferable to introduce at `this point the entire amount of makeup hydrogen to increase stripping efficiency. This fresh hydrogen is admitted via lines 80 and 75. If additional hydrogen is required, it is supplied via line 75 from separator 71. The stripped liquid-phase product is continuously withdrawn from stripper 77 via line 79, and transferred to storage or nal fractionation.

The liquid product in separator 71 is then transferred to low-pressure separator S1 via line S2 and valve 83. Final product is withdrawn via line 84 and oil-gases via line 85.

It will be understood that the temperatures, pressures, and feed rates in reactors 2, 30, 56 and 67 are substantially within the ranges above described as being generally applicable for hydrofining operations.

The catalysts employed herein may consist of any conventional hydrofining catalyst. In general, the oxides and -suliides of transitional metals are useful, and especially the group VI-B and group VIII metal oxides and suldes. In particular, the combination of one or more group VI-B metal oxide or sulde with one or more of group VIII metal oxide or sulde is preferred. For example, combinations of nickel-tungsten oxides and/ or suldes, cobalt-molybdenum oxides and/or suliides, are specifically contemplated. However, iron oxide, iron suliide, cobalt oxide, cobalt sulfide, nickel oxide, nickel suliide, chromium oxide, chromium suliide, molybdenum oxide, molybdenum sulide, tungsten oxide or tungsten sulfide may be used alone to less advantage.

In all the foregoing cases, it is preferable to support the active catalyst on a relatively inert carrier. Generally, minor proportions of the active metal compounds are used, ranging between about 1% `and 25% by weight. Suitable carriers include for example activated alumina, activated alumina-silica, Zirconia, titania, activated clays such as bauxite, bentonite and montmorillonite, and the like. Preferably the active components are added to the carrier by impregnation from aqueous solutions followed by drying and calcining to activate the composition. Suitable calcining temperatures range between about 500 and 1200 C.

The preferred catalyst for use herein comprises the composition usually known as cobalt-molybdate, which actually may be a mixture of cobalt vand molybdenum oxides. This mixture is preferably distended upon activated alumina, or still more preferably activated alumina containing 1% to 15% of copreciptated silica gel. The ratio of cobalt to molybdenum may be -between 0.4 and 5.0, and the total proportion of active ingredients is preferably between about 7% and 22% by weight, comprising about l-7% of C00, and 6-15% of M003. Catalysts of this type may be prepared by coprecipitation of both components on the carrier as described in U.S. Patent No. 2,369,432, and No. 2,325,033, or by co-impregnation of both components on the carrier as described in U.S. Patent No. 2,486,361. Preferably however the catalyst is prepared by separate alternate impregnations as described in U.S. Paten-t No. 2,687,381.

The catalyst employed in the mixed-phase hydroiiner may be identical to that employed in the vapor-phase hydroner, or the two may be diiferent. It has been found that in using mixtures of group VIE-group VIII type catalysts, it is preferable to use compositions containing a higher total proportion of group VIII metal, and higher group VIII/ group VIB metal ratios, when processing liquid-phase hydrocarbons than when only vapor-phase material is being processed. In vapor-phase processing using cobalt molybdate for example, the degree of conversion obtainable under a given set of conditions appears to level olf when the total cobalt content reaches about 3%, the molybdenum content remaining constant at about 10%. However, when a liquid phase is present it is found that substantially increased conversions are obtainable at cobalt oxide contents ranging up to about 7% It is therefore preferred -to use catalysts containing about 6 %-15 of molybdenum oxide in either case, butto use those containing `about 1 %-3 .5% of cobalt oxide in the vapor-phase hydrofining, :and those containing about 3.0-7.0% of cobalt oxide when a liquid phase is present. The same preferred ratios hold Where other group VIII metals are substituted for cobalt, and other group VIB metals for molybdenum.

Typical results obtainable in practice of the present invention are illustrated by the following examples. These examples however should not be construed as limiting in scope.

Example I It is desired to hydrofine a gas oil feed which is high in both sulfur and nitrogen. 'Ihe light fraction from this feed is to be used as jet fuel and requires a high degree of nitrogen removal. The heavy fraction of the feed is to lbe used as feedstock to a fluid catalytic cracking unit, and as much as 0.15% nitrogen can be tolerated therein. In order fto achieve the desired hydroning of yboth fractions, a processing scheme similar to that shown in Figure 1 is used. The catalyst employed in both reactors 2 `and 30 is 3 cobalt loxide-9% molybdenum oxide composite supported on `a alumina-5% silica carrier. rI'he fresh stripping hydrogen used in stripping column 3 comprises 900 s.c.f. per barrel of total feed of 95 hydrogen. The details of operating conditions, -feed and product characteristics are as follows:

Reactor 2 Reactor 30 Feed:

Gravity, API 25. 0 38. Boiling range, F-- 400-800 10D-700 Sulfur, Wt. percent. 2. 5 0.03 Nitrogen, wt. percent. 0.3 0. 05 Product:

Gravity, API 24. 0 40. 0 Boiling range, F 60G-750 10G-650 Sulfur, Wt. percent. 0. 07 0.001 Nitrogen, wt. percen 0.08 0. 004 Volume percent oi original feed 42 61 Reactor Operating Conditions:

emp.:

In, F 70o 78o Out, F- 780 800 Pressure:

In, p.s.i.g. 1100 1080 Out, p.s.i.g. 1090 1070 Hydrogen to oil ratio, s.c.f./bbl 2000 3600 Hydrogen Consumption, s.c.f./bbl 700 200 Liquid hourly space velocity 1. 5 1. 5

It will thus be apparent that the product from reactor 2 is suitable in quality for catalytic cracking charge stocks, while the product yfrom reactor 30 is suitable for use as ya jet fuel. The procedure of this example is especially useful for ,treating oils having an initial boiling point between about 350-450" F., and an end boiling point between yabout 750-850 F.

Example II The procedure of Example I is repeated, except that the catalyst in reactor 2 is modified to contain 5.0% by weight lof cobalt oxide. The catalyst in reactor 30 is the same 3% cobalt oxide-9% molybdenum oxide catalyst used in Example I. The respective products obtained from reactors 2 and 30 are found to contain about 10% less sulfur and 15% less nitrogen. When this operating sequence is repeated using a 9% molybdenum oxide-5% cobalt oxide catalyst in reactor 30, and the 9% molybdenum oxide-3% cobalt oxide catalyst in reactor 2, substantially no improvement in desulfurization or denitrogenation is obtained over that shown in Example I.

Example III Y Y 7 use the naphtha Vfraction thereof as "feed to a platinum catalyst reforming unit, and to kuse the lheayyfraction for diesel fuel. For use as -a catalytic reformer feed, the naphatha fraction must contain less than 0.05% sulfur and less than 0.000l% nitrogen. To process this feed, it is introduced into a reactor similar to that shown at 56 in Figure 2, and processed under the vconditions described in Example I for reactor 2. The effluent is then partially cooled Yto a temperature of about '500 F., and introduced into a liquid-gas separator. The liquid phase is stripped countercurrently with about 750 s.c.f. of fresh hydrogen per barrel of initial feed, the strippingfhydrogen being introduced at about 200 F. The vapor phase resulting from the stripping operation is allowed to mingle with `the vapor-phase naphtha from the rst hydroning operation, and the mixture is then reheated `to about 780 F. and passed through a second hydrofiner at about 1000 p.s.i.g. and a liquid hourly space velocity of 2.0. The naphtha product recovered is found tokhave a boiling range of about 250-450 F.

' and contains about-0.075% sulfur and 0.000l% nitrogen andis thus suitable reforming stock.

The liquid product has a boiling range of about 400 to 600 F., and contains about 0.05% sulfur:and-`0.06% nitrogen. lThe initial feedstock contained 2.1% sul-fur and 0.25% nitrogen. Y

The procedure of this example is especially luseful for treating oils having an initial boiling pointY between about`250350 F., vand `an end boiling point between about y6'00-7 50 F.

Results similar to those described in the examples are obtained when other catalysts ,within the scope of this invention are substituted for the cobalt-molybdate. Simi- -larly, other processing conditions may beemployed to obtain commensurate benefits. The scope of the invention should not be construed as limited to the exemplary details. The true scope of the invention is intended to be embraced by the following claims.

We claim:

'1. In a process wherein a mineral'oil feedstock having an end-boiling-point above about 500 F. and an :initialboiling-point at least 75 F. low-er than said end-boilingpoint is subjected to mixed-phase hydrofining at an elevated pressure inthe presence of hydrogen and `a catalyst, and wherein the conditions of hydrofining are such that a portion of said feedstock is in the vapor-phase and another portion is in the liquid-phase, the improvement which comprises (Y1) terminating said mixed-phase contacting after said liquid-phase'has been sufficiently treated to decompose a substantial desired proportion of organic impurities therein but before said vapor-phase has been sufficiently treated to eect removal therefrom of the desired proportion of organic impurities, (2) separating the total vapor-phase product from the liquid-phase product at substantially the pressure prevailing in said mixedphase hydroiining, (3) subjecting said liquid-phase product to countercurrent stripping with stream at substantially the pressure prevailing in said mixed-phase hydroiining to strip dissolved low boiling hydrocarbons therefrom, (4) mixing said separated vapor-phase product with the vapor-phase stripping efiluent from saidstripping step, yand (5) subjecting the resulting combined vapor-phase mixtureito further hydrofming at an'elevated pressure in the presence of a hydroiining catalyst but substantially in the absence of a liquid phase to effect further removal of organic impurities therefrom.

2. A process as defined in claim 1 wherein the prod- ,ucts from said vapor phase hydrofining are .cooled Iand condensed Without substantial reduction in pressure, and the resulting hydrogen-rich gas phase is recycled to said mixed-phase hydroning. g

3'. A process as defined in claim l wherein the catalysts `.employed -forrsaid mixed-phase hydroning and said vaporgphase hydroning comprise two active components 'in intimate admixture, one of said components being sea hydrogen-rich gas lected from the class consisting ofthe oxides and 'suldes of group VIB metals, and therother of said components being selected from the class consisting of the oxides and suliides of group VIII metals. Y 1

4. A process as defined in claim l wherein the catalysts employed in said mixed-phase hydroiining and said vaporphase hydroning consist essentially of cobalt molybdate distended on a major proportion of an adsorbent oxide carrier.

5 A process as Vdefined in claim l wherein the catalyst employed in said vapor-phase hydroiining is essentially cobalt molybdate'onralumina wherein the cobalt oxide content is between about 1% .and 3.5% by weight, `and the catalyst employed in vsaid mixed-phase hydroiining is essentially cobalt molybdateon alumina whereinthe cobalt oxide content is (l) `higher than that of said vaporphasehydroiining catalyst, and (2) between about 3% and 7% by weight, the molybdenum oxide content of each of said catalysts beingbetween about 6% and `15% by weight.

6. A process `as dened in claim 1 wherein the total effluent from said mixed-phase 'hydroning is partially cooled to eiect a partial condensation of the vapor phase Yprior to said separation of vapor phase from liquid phase.

7. A method for hydroining `a gas toil feedstock having an initial boilingpoint between about 350-450 F., and an end boiling point between about 750-850 F., to obtain therefrom a light fraction suitable for use as jet fuel, and a heavy fraction suitable for use as catalytic cracking charge stock, which comprises first subjecting said feedstock to mixed-phase hydrofining at an elevated pressureY in the presence of hydrogen and a catalyst, wherein the conditions of hydrofining are such that the heavy ends are predominantly inthe liquid phase and the light ends are predominantly in the vapor phase, both of said phases being in concurrent downiiow while in contact with said catalyst, terminating said mixed-phase contacting after said mixed-phase hydroiining, and subjecting the resulting combined vapor-phase mixture to further hydroining `at an elevated pressure in the presence of -a hydroning catalyst and in the absence of liquid phase, recovering from said stripping operation a heavy fraction suitable for use as cracking charge stock, .and recovering from said vapor-phase hydroiining a light fraction suitable for use as jet fuel.

8. A process as defined in claim 7 wherein the total eiiuent from said mixed-phase hydroiinjng is partially cooled to effect ya partial condensation of the vapor phase prior to said separation of vapor phase from liquid phase.

9. A process as defined in claim 7 wherein the catalysts employed in said mixed-phase hydrotining and said Vaporphase hydroiining consist essentially of cobalt molybdate ldistended on a major proportion of an adsorbent oxide carrier.

l0. A method for hydroning a gas oil feedstock having an initial boiling point between about Z50-350 F. and an end boiling point between about 600-7507 F. to obtain therefrom a naphtha fraction suitable for reforming charge stock, and a heavy fraction suitable for Yuse as diesel fuel, which comprises` first subjecting said are predominantly in the vapor phase, both of said phases being in concurrent downflow while in contact with said catalyst, terminating said mixed-phase 4cont-acting after said liquid phase has been su'iciently rened for use as diesel fuel but before said vapor phase has been sufciently refined for use as reforming charge stock, separating the total vapor-phase product from the liquidphase product at substantially the pressure prevailing in said mixedphase hydroning, stripping said liquid phase with fresh hydrogen at substantially the pressure prevailing in said mixed-phase hydrotining to strip dissolved low-boiling hydrocarbons therefrom, mixing the vapor phase resulting from said stripping with the vaporphase product from said mixed-phase hydroning, and subjecting the resulting combined vapor-phase mixture to further hydroiining at an elevated pressure in the presence of a hydrotining catalyst and in the absence of liquid phase, recovering from said stripping operation a heavy fraction suitable for use as diesel fuel and 10 recovering from said vapor-phase hydrotining a naphtha fraction suitable for use as reforming charge stock.

11. A process as dened in claim 10 wherein the total euent from said mixed-phase hydroning is partially cooled to effect a partial condensation of the vapor phase prior to said separation of vapor phase IJfrom liquid phase.

12. A process as dened in claim 10 wherein the catalysts employed in said mixed-phase hydroning and said vapor-phase hydroning consist essentially of cobalt molybdate distended on a major proportion of an adsorbent oxide carrier.

References Cited in the tile of this patent UNITED STATES PATENTS 1,908,286 Dorrer May 9, 1933 2,587,987 Franklin Mar. 4, 1952 2,769,354 Sweetser et al. Nov. 6, 1956 

1. IN A PROCESS WHEREIN A MINERAL OIL FEEDSTOCK HAVING AN END-BOILING-POINT ABOVE ABOUT 500*F. AND AN INITIALBOILING-POINT AT LEAST 75*F. LOWER THAN SAID END-BOILINGPOINT IS SUBJECTED TO MIXED-PHASE HYDROFINING AT AN ELEVATED PRESSURE IN THE PRESENCE OF HYDROGEN AND A CATALYST, AND WHEREIN THE CONDITIONS OF HYDROFINING ARE SUCH THAT A PORTION OF SAID FEEDSTOCK IIS IN THE VAPOR-PHASE AND ANOTHER PORTION IS IN THE LIQUID-PHASE, AND IMPROVEMENT WHICH COMPRISES (1) TERMINATING SIAD MIXED-PHASE CONTACTING AFTER SAID LIQUID-PHASE HAS BEEN SUFFICIENTLY TESTED TO DECOMPOSE A SUBSTANTIAL DESIRED PROPORTION OF ORGANIC IMPURITIES THEREIN BUT BEFORE SAID VAPOR-PHASE HAS BEEN SUFFICIENTLY TREATED TO EFFECT REMOVAL THEREFROM OF THE DESIRED PROPORTION OF ORGANIC IMPURITIES, (2) SEPARATING THE TOTAL VAPOR-PHASE PRODUCT FROM THE LIQUID-PHASE PRODUCT 